The present invention relates to an improved circulating, i.e., fast, fluidized bed reactor having two stages, namely, a circulating fluidized bed reaction stage and a cyclonic reaction stage downstream of the fluidized bed; and to a method of operating the reactor. More particularly, the invention relates to a two stage circulating fluidized bed reactor in which the size of the fluidized bed reaction chamber and the cyclonic reaction vessel are substantially reduced.
The present invention has specific application, inter alia, to adiabatic fluidized bed combustors, fluidized bed boilers, and compressed hot air generators. As used herein, and in the accompanying claims, "adiabatic combustor" denotes a fluidized bed combustor that does not contain internal cooling means, and "boiler" denotes a fluidized bed combustor that contains internal heat absorption means, in the form of boiler, superheater, evaporator, and/or economizer heat exchange surfaces. The temperature of adiabatic fluidized bed combustors is typically controlled by the use of pressurized air in substantial excess of the stoichiometric amount needed for combustion. On the other hand, fluidized bed boilers require very low excess air, so that heat absorption means are required in the fluidized bed. Fluidized bed gasifiers, in contrast, utilize less than stoichiometric amounts of air.
The state of fluidization in a fluidized bed of solid particles is primarily dependent upon the diameter of the particles and the fluidizing gas velocity. At relatively low fluidizing gas velocities exceeding the minimum fluidizing velocity, the bed of particles is in what has been termed the "bubbling" regime. Historically, the term "fluidized bed" has denoted operation in the bubbling regime. This fluidization mode is generally characterized by a relatively dense bed having an essentially distinct upper bed surface, with little entrainment, or carryover, of the bed particles (solids) in the flue gas, so that recycling the solids is generally unnecessary. At higher fluidizing gas velocities, above those of the bubbling regime, the upper surface of the bed becomes progressively diffuse and carry-over of the solids increases, so that recirculation of solids using a particulate separator, e.g., a cyclone separator, becomes necessary in order to preserve a constant solids inventory in the bed.
The amount of solids carry-over depends upon the fluidizing gas velocity and the distance above the bed at which the carry-over occurs. If this distance is above the transfer disengaging height, carry-over is maintained at a constant level, as if the fluidizing gas were "saturated" with solids.
If the fluidizing gas velocity is increased above that of the bubbling regime, the bed then enters what has been termed the "turbulent" regime, and finally, the "fast," i.e., "circulating" regime. If a given solids inventory is maintained in the bed, and the fluidizing gas velocity is increased just above that of the turbulent regime, the bed density drops sharply over a narrow velocity range. Obviously, if a constant solids inventory is to be preserved in the bed, the recirculation, or return, of solids must equal the carry-over at "saturation."
At fluidizing gas velocities below those associated with the aforementioned sharp drop in bed density, the effect upon bed density of returning solids to the fluidized bed at a rate well above the "saturation" carry-over is not marked. The addition of solids to a bed fluidized in either the bubbling or turbulent regime at a rate above the saturation carry-over will simply cause the vessel containing the fluidized bed to fill up continually, while the fluidized density will remain substantially constant. However, at the higher fluidizing gas velocities associated with the circulating regime, the fluidized density becomes a marked function of the solids recirculation rate.
Circulating fluidized beds afford intimate contact between the high velocity fluidizing gas and large inventory of solids surface per unit bed volume. Additionally, slip velocity (i.e., solids-fluidizing gas relative velocity) is relatively high in circulating fluidized beds, when compared with that in ordinary fluidized beds. Consequently, there is generally a very high level of particulate loading in the combustion gases exiting from circulating fluidized bed combustors. The combustion process which takes place in a circulating fluidized bed combustor is also generally more intense, having a higher combustion rate than that occurring in traditional fluidized bed combustors. Furthermore, as a result of the high solids recirculation rate in circulating fluidized beds, the temperature is essentially uniform over the entire height of such combustors.
Conventional circulating fluidized bed combustors operate at gas superficial velocities many times higher than the terminal velocity of the fluidized bed mean particle. Consequently, there is a very high particulate loading in the combustion product gases exiting from the combustor and entering the downstream cyclone particle separator. Such conventional cyclone particle separators typically have a height which is roughly three times their diameter, so that separators having a large diameter designed to remove the entrained solids from circulating fluidized bed combustors are typically quite tall and bulky. Such large refractory coned cyclone particle separators constitute a significant portion of the total cost of conventional circulating fluidized bed combustor systems.
Notwithstanding the many advantages offered by conventional circulating fluidized bed reactors, as enumerated above, the high cost of constructing and maintaining the extremely large cyclone particle (gas-solids) separators required for recirculation of the entrained solids at the rate necessary to maintain the bed in the circulating fluidization regime constitutes a severe economic impediment to widespread commercial utilization of such reactors.
Prior art circulating fluidized bed combustor boilers are known which employ vertical heat exchanger tube-lined walls in the entrainment region of the combustor (i.e., parallel to the flow). Such combustors rely primarily on the transfer of heat from gases which typically are heavily laden with solids, and require an extremely large internal volume to accomodate the large heat transfer surface required.
The tube-lined wall heat transfer surface installed in the free board region in conventional fluidized bed combustors necessarily possesses a significantly lower heat transfer coefficient than that of a heat transfer fully immersed in the fluidized bed. Furthermore, its heat transfer coefficient is dependent primarily on two parameters: (a) fluidizing gas velocity, and (b) particle concentration in the flue gases, i.e., particle loading. The latter parameter is, in turn, strongly dependent on the fluidizing gas velocity and the mean particle size of the fluidized bed material. The concentration of particles in the ascending gas flow in a conventional circulating fluidized bed combustor is directly proportional to the gas velocity to the 3.5-4.5 power, approximately, and inversely proportional to the fluidized bed mean particle diameter to the 3.0 power, approximately. The strong effect of, and careful attention to, these two parameters on the concentration of particles in the ascending gas flow helps to achieve a reasonable heat transfer coefficient for conventional tube-lined wall heat transfer surfaces in the free board region and facilitates the control of combustion temperature at nominal and reduced boiler capacity. Nevertheless, there is a need in the art for a fluidized bed combustor boiler having a reasonable heat transfer coefficient and permitting control of combustion temperature at nominal and reduced capacity without being so strongly dependent on fluidizing gas velocity and fluidized bed mean particle diameter.
The height of the free board region of a conventional circulating fluidized bed combustor boiler having a tube-lined wall heat transfer surface as described above is directly proportional to the superficial gas velocity to the 0.5 power and inversely proportional to the surface's heat transfer coefficient. Also, it can be shown that the particle loading and heat transfer coefficient are directly proportional to any change in the superficial gas velocity. The latter fact means that, for instance, a reduction of the superficial gas velocity will require an incease in the free board height for such a conventional combustor of a given capacity. Similarly, it can be shown that in order to increase the capacity of such a combustor, the free board height must be increased, thereby significantly increasing the cost of constructing such a higher capacity combustor.
In contrast to most conventional circulating fluidized bed combustors, the combustor disclosed in U.S. Pat. No. 4,469,050 to Korenberg (assigned to a common assignee herewith) does not provide for transferring the entrained granular bed material, unburnt fuel, ash, gases, etc. directly into a cyclone particle separator. Rather, the entrained solids and gases are carried upward into a cylindrically shaped upper region of the combustor chamber, i.e., an extended free board region, where further combustion takes place. Vertical rows of tangential nozzles are built into and evenly spaced over this cylindrical upper free board region. This tangentially fed secondary air is supplied at a sufficient velocity, and the geometric characteristics of the cylindrical upper region are adapted, to provide a Swirl number (S) of at least about 0.6 and a Reynolds number (Re) of at least about 18,000 within such upper region, which are required to create a cyclone of turbulence.
This cyclone of turbulence enables the combustor shown in U.S. Pat. No. 4,469,050 to achieve specific heat releases higher than 1.5 million Kcal per cubic meter per hour, thereby significantly increasing the rate of combustion. As a direct result, the "vessel" size of this combustor is significantly smaller than other prior art combustors. Essentially, compared to its downstream cyclone particle separator, the combustor vessel appears like a refractory-lined duct.
The relatively large size of the cyclone particle separator compared to the combustor vessel produced an incentive for improving this system by eliminating the cyclone particle separator. This was achieved in the circulating fluidized bed combustor disclosed in U.S. Pat. No. 4,457,289 to Korenberg (assigned to a common assignee herewith) by eliminating the entire external solids recirculation loop and utilizing "internal recirculation." To achieve this, a "throat" was inserted at the top of the cylindrical upper region of the combustor and the external cyclone separator was eliminated.
The combustor disclosed in U.S. Pat. No. 4,457,289 is significantly less expensive to construct than the combustor disclosed in U.S. Pat. No. 4,469,050, and other prior art circulating fluidized bed combustors, since it does not require a separate cyclone particle separator. However, it has demonstrated a somewhat reduced particulate capturing efficiency compared to such other combustors, particularly when burning solid coal particles. Furthermore, the combustor disclosed in U.S. Pat. No., 4,457,289 provides a residence time for solid coal particles and conventional sulfur absorbents which, in some cases, may be less than optimum for capturing any sulfur in the coal.
In conventional, non-circulating and circulating fluidized bed reactors for combusting particulate material, the material to be combusted is fed in or over a bed of granular material, usually fuel ash, sulfur absorbents such as limestone, and/or sand.